US 7938887 B2
Compositions and methods related to the removal of acidic gas. In particular, the present disclosure relates to a composition and method for the removal of acidic gas from a gas mixture using a solvent comprising a diamine (e.g., piperazine) and carbon dioxide. One example of a method may involve a method for removing acidic gas comprising contacting a gas mixture having an acidic gas with a solvent, wherein the solvent comprises piperazine in an amount of from about 4 to about 20 moles/kg of water, and carbon dioxide in an amount of from about 0.3 to about 0.9 moles per mole of piperazine.
1. An aqueous composition comprising:
a diamine in an amount of from about 4 to about 20 moles/kg of water; and
carbon dioxide in an amount of from about 0.3 to about 0.9 moles carbon dioxide per mole of the diamine.
2. The composition of
3. The composition of
4. The composition of
5. The composition of
6. A method for reducing the volatility of a diamine comprising:
providing a solvent that comprises a diamine in an amount of from about 4 to about 20 moles/kg of water; and
adding carbon dioxide to the solvent in an amount of from about 0.3 to about 0.9 moles per mole of diamine.
7. The method of
8. The method of
9. The method of
10. A method for increasing the solubility of a solid diamine comprising providing a diamine to a solution of from about 0.3 to about 0.9 moles carbon dioxide per mole of the diamine.
11. The method of
12. The method of
13. A method comprising contacting a gas mixture comprising an acidic gas with a solvent, wherein the solvent comprises a diamine in an amount of from about 4 to about 20 moles/kg of water.
14. The method of
15. The method of
16. The method of
17. A method comprising:
contacting a gas mixture comprising an acidic gas with a solvent, wherein the solvent comprises a diamine in an amount of from about 4 to about 20 moles/kg of water; and
allowing the acidic gas to transfer from the gas mixture to the solvent.
18. The method of
19. The method of
20. The method of
21. The method of
22. The method of
23. The method of
24. The method of
25. The method of
26. The method of
27. A method comprising:
contacting a gas mixture comprising an acidic gas with a solvent in an absorber, wherein the solvent comprises piperazine or a substituted piperazine in an amount of from about 4 to about 20 moles/kg of water;
allowing the acidic gas to transfer from the gas mixture to the solvent;
forming a purified gaseous stream and a rich solvent stream; and
routing the rich solvent stream through a multistage flash stripper, wherein the multistage flash stripper is operated at temperatures less than about 175° C.
This application claims the benefit of U.S. Provisional Application Ser. No. 61/019,646, filed Jan. 8, 2008, which is incorporated in this application by reference.
The present invention was made with support under contract number FC26-2NT41440 awarded by the Department of Energy. The U.S. government has certain rights in the invention.
As concerns of global climate changes spark initiatives to reduce carbon dioxide emissions, its economic removal from gas streams is becoming increasingly important. Removal by absorption/stripping is a commercially promising technology, as it is well suited to sequester carbon dioxide (CO2). Such carbon dioxide emissions may be produced by a variety of different processes, such as the gas stream produced by coal-fired power plants. The removal of CO2 can be an expensive process, potentially increasing the cost of electricity by 50% or more. Therefore, technology improvements to reduce the costs associated with the removal are highly desirable.
The removal of CO2 from fuel gas and flue gas by absorption/stripping with aqueous amines is a disclosed and commercially practiced technology. A typical flowsheet for such a process is give by Kohl and Nielsen (1997) (
Commercially used amines that are used by themselves in water include monoethanolamine, diethanolamine, methyldiethanolamine, diglycolamine, diisopropanolamine, some hindered amines, and others (Kohl and Nielsen (1997)). These amines are soluble or miscible with water at ambient temperature at high concentrations that are used in the process to maximize capacity and reduce sensible heat requirements. Other amines, including piperazine, are used in combination with methyldiethanolamine and other primary amines.
A number of mono- and polyamines, including piperazine, are identified as potentially useful solvent components but have not been used because they are insufficiently soluble in water when used by themselves. Piperazine is a diamine that has previously been studied as a promoter for amine systems to improve kinetics. In water at 25° C., solid piperazine has a solubility less than 2 M, so it cannot be used in traditional systems at concentrations that give adequate CO2 capacity for good energy performance. BASF has disclosed the used of piperazine in combination with other amines (such as alkanolamines) or highly water soluble organics (such as triethyleneglycol) to promote the water solubility of piperazine.
It has also been claimed that number of potentially useful amines such as piperazine would be too volatile if used in high concentrations in aqueous solvents. The boiling point of piperazine (146.5° C.) is lower that that of monoethanolamine (170° C.), so the use of Raoult's law would suggest that it would have a greater volatility at the top of the absorber.
Some specific example embodiments of the disclosure may be understood by referring, in part, to the following description and the accompanying drawings.
The patent or application file contains at least one drawing executed in color. Copies of this patent or patent application publication with color drawing(s) will be provided by the Office upon request and payment of the necessary fee.
While the present disclosure is susceptible to various modifications and alternative forms, specific example embodiments have been shown in the figures and are described in more detail below. It should be understood, however, that the description of specific example embodiments is not intended to limit the invention to the particular forms disclosed, but on the contrary, this disclosure is to cover all modifications and equivalents as illustrated, in part, by the appended claims.
The present disclosure, according to certain embodiments, generally relates to compositions, systems, and methods for the removal of acidic gas. In particular, the present disclosure relates to compositions, systems, and methods for the removal of acidic gas from a gas mixture using a solvent comprising a diamine (e.g., piperazine) and carbon dioxide. Suitable diamines include those that are volatile or insoluble at normal conditions, for example, piperazine, ethylenediamine, and various substituted ethylenediamines (e.g., methylethylenediamine, dimethylethylenediamine, ethylethylenediamine, and diethylethylenediamine), and piperazine (e.g., methylpiperazine, dimethylpiperazine, ethylpiperazine, and diethylpiperazine).
The present disclosure is based in part on the discovery of optimum stripper process configurations and operating conditions that result in unexpectedly high lean loading of CO2. Such process configurations may include the matrix, internal exchange, flashing feed, and multipressure processes described herein. Such operating conditions may include an unexpectedly low exchanger approach temperature. In some embodiments, such exchanger approach temperatures may be approximately 5° C.
The present disclosure is also based in part on the discovery that diamines (e.g., piperazine) may be less volatile in an aqueous solution than expected from Raoult's law. In certain embodiments, the activity coefficient of piperazine at infinite dilution in water may be about 0.05, whereas monoethanolamine (MEA) has an activity coefficient of about 0.16.
The present disclosure is also based in part on the discovery that when diamine (e.g., piperazine) solutions are loaded with up to about 0.8 moles of carbon dioxide (CO2) per mole of diamine, the volatility of the diamine may be further reduced. Such a reduction may occur at least in part because of the formation of carbamate ions. Such a reduction may result in the ability to produce concentrated solutions of diamine loaded with CO2 which have a volatility acceptable for use in the methods of the present disclosure.
The present disclosure is also based in part on the discovery that the total solubility of solid diamine (e.g., piperazine) may be enhanced in solutions loaded with CO2. In certain embodiments, the present disclosure provides solutions comprising up to about 10 m (moles diamine/kg water) total diamine when said solutions are loaded with about 0.8 moles CO2 per mole of diamine. This increase in solubility may be due in part to the formation of carbamate ions.
The present disclosure is also based in part on the discovery that concentrated aqueous diamine (e.g., piperazine) may be more stable to oxidative and/or thermal degradation as compared to conventional solutions, such as MEA. In certain embodiments, the presence of dissolved iron may catalyze the degradation of MEA at a higher rate than the degradation of diamine. In certain embodiments, solutions of diamine loaded with CO2 may not degrade significantly even at temperatures as high as 150° C., whereas MEA may undergo significant degradation (up to about 50%) at 120° C. Thus, in certain embodiments, the present disclosure provides solutions comprising a diamine which may be used advantageously at higher pressures and/or temperatures. For example, the solutions comprising a diamine may be used at temperatures less than 175° C. Such an ability to operate at higher pressures and/or temperatures may, among other things, reduce the amount of energy necessary to perform the methods of the present disclosure. In certain embodiments, such a reduction of the amount of energy may range from about 10% to about 30%. Additionally, solutions comprising diamine (e.g., piperazine) may absorb CO2 at faster rates. In certain embodiments, the use of solutions comprising piperazine may result in increased in CO2 absorption rates ranging from about 20% to about 100%. Such increased CO2 absorption rates may, among other things, enable absorber configurations which require less packing and pressure drop.
When used in the methods of the present invention, the diamine (e.g., piperazine) may be recovered following absorption of CO2. In certain embodiments, such recovery may occur through an evaporation process using a thermal reclaimer.
In certain embodiments, the present disclosure provides a method for the removal of acidic gases from a gas mixture comprising contacting the gas mixture with a solvent comprising a diamine (e.g., piperazine) in an amount of from about 4 to about 15 moles/kg of water and carbon dioxide in an amount of from about 0.3 to about 0.8 moles per mole of the diamine.
While the present disclosure primarily discusses removal of CO2, any acidic gas capable of removal by the methods of the present invention is contemplated by the present disclosure. Such acidic gases may include, but are not limited to, hydrogen sulfide (H2S) or carbonyl sulfide (COS), CS2, and mercaptans.
The gas mixture may be any gas mixture comprising CO2 for which CO2 removal is desired and which is compatible with (i.e. will not be adversely affected by, or will not adversely react with) the methods of the present disclosure. In certain embodiments, the gas mixture may comprise any gas mixture produced as the byproduct of a chemical process. Suitable gas mixtures may comprise one or more of natural gas and hydrogen.
In certain embodiments, the present disclosure provides several process configurations that may be useful in the methods of the present disclosure. The choice of process configuration may depend upon a number of factors, including, but not limited to, the composition of the gas mixture, the desired amount of CO2 removal, the concentration of diamine (e.g., piperazine) to be used, and resource or environmental considerations.
One type of process configuration that may be useful in the methods of the present invention is a matrix stripper configuration. In certain embodiments, such a matrix stripper configuration may be a two-stage matrix, such as the configuration shown in
Another type of process configuration that may be useful in the methods of the present invention is multistage flash stripper configuration. In certain embodiments, such multistage flash strippers may be configured as a multistage flash with a multistage intercooled compressor as illustrated in
Another type of process configuration that may be useful in the methods of the present invention is an exchange stripper configuration. In certain embodiments, such an exchange stripper configuration may be an internal exchange stripper, such as the configuration shown in
Another type of process configuration that may be useful in the methods of the present invention is a multipressure configuration. In certain embodiments, such a multipressure configuration may be a multipressure configuration with a split feed, such as the configuration shown in
Another type of process configuration that may be useful in the methods of the present invention is a flashing feed configuration. An example of such a configuration is shown in
The choice of operating conditions for each of these process configurations may depend upon a number of factors, including, but not limited to, the composition of the gas mixture, the desired amount of CO2 removal, the concentration of piperazine to be used, and resource or environmental considerations. Examples of suitable operating conditions are shown in
Therefore, the present invention is well adapted to attain the ends and advantages mentioned as well as those that are inherent therein. While numerous changes may be made by those skilled in the art, such changes are encompassed within the spirit of this invention as illustrated, in part, by the appended claims.
Solubility of piperazine (PZ.6H2O) solid in PZ solvents. The solubility of PZ.6H2O in water was measured by differential scanning calorimetry. The vapor pressure of PZ and CO2 was measured over loaded and unloaded solutions of piperazine from 30 to 60° C. These data were combined with other thermodynamic data on PZ/CO2/H2O and regressed with the electrolyte/NRTL model in AspenPlus.
Loaded solutions of piperazine were prepared by mixing and heating anhydrous piperazine and water, then sparging with CO2 on a balance. Table 1 shows that solutions with loading less than 0.3 and greater than 0.4 usually precipitate solids when cooled to ambient temperature. The solids precipitated at greater loading were identified as hydrated, protonated piperazine carbamate (CO2.piperazine.H2O).
The solid solubility of PZ was studied over a range of PZ concentration, CO2 loading, and temperature. Solutions were prepared to cover the desired solution properties and were allowed to equilibrate at each condition with stirring before solubility observations were made. The transition temperature of 8 and 10 m PZ solutions over a range of CO2 loading is shown in
Viscosity. The viscosity of aqueous PZ solutions has been measured from 0.20 to 0.45 moles CO2/mole alkalinity, 2 m PZ to 20 m PZ, and 25° C. to 60° C. The viscosity of 8 and 10 m PZ is compared with other amines in
Comparison of the viscosity on this basis shows how the amine basic groups affect overall viscosity. As the concentration of basic groups increases, the viscosity increases in a linear direction. The viscosity of 8 m PZ is higher than that of 7 m MEA, but as compared to 60 wt % DGA®, the viscosity of PZ is lower for a higher alkalinity. DGA® solutions at 60 wt % are successfully used in natural gas treating.
Oxidative Degradation. Heavy metals are known to catalyze the oxidative degradation of amines. The results of oxidative degradation of concentrated PZ in the presence of several dissolved metals are shown in Table 2. The experiments simulated four scenarios: (1) leaching of stainless steel metals, (2) addition of a copper-based corrosion inhibitor, (3) addition of a vanadium-based corrosion inhibitor (low concentration), and (4) addition of a copper-based corrosion inhibitor and proprietary inhibitor “A”.
Oxidative degradation of concentrated PZ was found to be four times slower than that of MEA in the presence of stainless steel metals (Fe2+, Cr3+, and Ni2+) and a low concentration of vanadium. As with MEA solutions, PZ was determined to be highly susceptible to oxidative degradation in the presence of Cu2+. The primary degradation products were found to be ethylenediamine (EDA), formate, oxalate, and N-formylpiperazine, the amide of formate and PZ (denoted as Formamide in the table). The N-formylpiperazine concentration was not measured directly, but inferred from formate production through the basic reversal of the N-formylpiperazine formation reaction. Also, as with MEA, Inhibitor “A” was able to vastly reduce this degradation to levels comparable with the stainless steel and vanadium cases.
Thermal Degradation. Thermal degradation was investigated in PZ solutions at slightly above stripper temperature (135° C.) and much higher than stripper temperatures (150° C. and 175° C.). The thermal degradation results are shown in Table 3. Experiments ranged from 4 to 12 weeks in length.
PZ thermal degradation was determined to be negligible at 135 and 150° C. as compared to 7 m MEA. At 175° C., 32% of the PZ was degraded in 4 weeks. EDA was observed as a thermal degradation product at 175° C. but not at lower temperatures. Addition of 5.0 mM Cu2+/0.1 mM Fe2+, 5.0 mM Cu2+/0.1 mM Fe2+/100 mM Inhibitor “A”, and 0.6 mM Cr3+/0.25 mM Fe2+/0.25 mM Ni2+ did not affect degradation rates at 175° C.
The amine concentration was determined by cation chromatography.
CO2 Solubility. The measured solubility of CO2 in 2 m to 8 m PZ solutions is in given in
Kinetics of CO2 Absorption in PZ Solutions. The kinetics of the CO2 absorption into concentrated aqueous PZ was studied in a wetted wall column. The measured liquid-side mass transfer coefficient based on a gas side driving force, kg′, is shown compared to 7 m MEA in
As demonstrated in
Volatility of PZ Solutions. The volatility of PZ was measured in the equilibrium cell with hot gas FTIR. The volatility of 8 m PZ solutions is compared to that of 5 m PZ and 7 m MEA in
At 40° C., the normalized volatility of PZ solutions is the same as the normalized volatility of MEA solutions. It was anticipated that PZ would have a higher volatility than MEA because the boiling point of PZ, 146° C., is lower than that of MEA, 170° C. However, the volatility of PZ is comparable at 40° C. Initial modeling of PZ systems demonstrates this effect as a greatly reduced activity coefficient for PZ. At 40° C., PZ volatility varies from 10 to 19 ppm at atmospheric pressure.
Cullinane (2005) estimated PZ volatility over lean solutions (
Kohl and Nielsen (1997) note that “CO2 stripping from MEA solution increases with increased reboiler pressure.” Elevated stripper pressure/temperature has also been identified as a useful operating condition to minimize total energy requirement (PCT/US2004/019838). This effect results from temperature swing desorption, because the heat of CO2 desorption is greater that the heat of H2O evaporation, less heat is required to strip CO2 at greater stripper temperature. This is especially true with solvents that have a greater heat of CO2 desorption.
The heat of CO2 desorption was estimated by the rigorous thermodynamic model based on data with 2.4 m PZ. Some results are given in Table 4.
In PZ solutions the heat of CO2 desorption is always greater than that of water, which is about 40 kJ/gmol. Therefore, it may be advantageous to operate the stripper at greater temperature. Because PZ does not degrade very fast, even at greater temperature it should be possible to operate the stripper as hot as 135 to 160° C., giving improved energy performance.
Estimated Energy Requirement. Thermodynamic models for MEA and PZ were developed by Hilliard, and the PZ model was modified for concentrated solutions. The stripper section of an absorber/stripper system for CO2 removal was simulated for 8 m PZ and compared with 7 m MEA. These simulations included a simple stripper with CO2 compression to 5 MPa, a 5° C. cold side temperature approach for the cross heat exchanger, and a 10° C. approach for the reboiler. For all cases, 15 m of CMR NO-2P packing was used with an 80% approach to flood. The rich stream for each case assumed a P*CO2 at the absorber temperature of 40° C. One PZ case assumed a higher P*CO2 due to the faster rates of PZ expected in the absorber.
Each system was simulated at their optimum lean loadings and the baseline system, 7 m MEA, had an equivalent work of 36.1 kJ/mol CO2. The two 8 m PZ systems modeled at a 5.0 or 7.5 kPa rich equilibrium CO2 partial pressure had minimum equivalent works of 33.5 kJ/mol CO2 and 32.6 kJ/mol CO2, respectively. The PZ system with the lower rich P*CO2, 5 kPa, was less efficient than the system with 7.5 kPa P*CO2, but was better than the 7 m MEA case with an equivalent rich loading. The increased capacity of PZ improved its performance over the baseline, despite a lower ΔHabs.
Testing of Process Configurations.
Model Development. An equilibrium stripper model for aqueous solvents developed in Aspen Custom Modeler (ACM) was used to evaluate the different process configurations and solvents. A rich end pinch is usually predicted because of the generous amount of contacting assumed in the model. The stripper consisted of a flash region, 10 segments with 40% Murphree efficiency assigned to CO2, and a reboiler with 100% CO2 efficiency. The flash region in the column was quantified in terms of actual section performance.
Modeling Assumptions. The following modeling assumptions were made:
The CO2 vapor pressure (kPa) under stripper conditions for 7 m MEA, promoted MEA and different PZ/K2CO3 blends is represented by the empirical expression in
The CO2 vapor pressure over 4.28M MDEA and KS-1 (a proprietary solvent described by Mimura et al., based on the model by Posey et al. is shown in
The heat of desorption for 4.28M MDEA and KS-1 was assumed to be constant at 62 and 73 kJ/gmol CO2, respectively. The heat of vaporization of water, partial pressure of water, and heat capacities of solvent (assumed to be water), steam, and CO2 were calculated with equations from the DIPPR database as described by Fisher et al. The molar heat capacities for the CO2, water and amine were assumed to be equal and set to that of one mole of water.
The partial pressure of CO2 and water in each section was calculated by the following equation:
The model inputs were the rich loading and liquid rate, the temperature approach on the hot side of the cross exchanger (difference between the temperature of the rich stripper feed and the lean solution leaving the bottom of the stripper), and column pressure. Initial guesses of the lean loading, section temperatures, partial pressures, and loading were provided. The model solves equations for calculating VLE, and for material and energy balances. It calculates temperature and composition profiles, reboiler duty, and equivalent work.
The total energy required by the stripper is given as total equivalent work:
The work lost by extracting steam from the power plant, which would have been used to drive turbines to generate electricity, is the first term on the righthand side of the equation in the previous paragraph, while the second is the compressor work. The condensing temperature of the steam is assumed to be 10 K higher than the reboiler fluid. The turbine assumes condensing steam at 313 K, and has been assigned an effective efficiency of 75%.
In this example, the lean loading for each configuration was optimized to minimize equivalent work. The optimum lean loading, the lean loading that minimized equivalent work, was quite flat for cases with ΔT=5 K, and was approximately that for 90% change in equilibrium partial pressure of CO2 from the absorber rich end to the absorber lean end at 313 K. The cases with an approach temperature of 10 K frequently resulted in greater that 90% change in equilibrium partial pressure (overstripping) from the absorber rich to lean ends.
In all of the cases with a 5 K approach temperature, the optimum lean loading was quite high. The heat of absorption shown in
Effect of Varying Temperature Approach. The ‘baseline’ configuration for each system given in
Effect of Operating Pressure. Operating the stripper atrocuum (30 kPa), with a 5 K temperature approach in the cross exchanger, offers a 14% reduction in equivalent work for 6.4 m K1/1.6 m PZ and 4 and 20% more energy with 5 m K1/2.5 m PZ and MEA/PZ, respectively. This shows that vacuum operation favors solvents with low heats of absorption, while operation at normal pressure favors solvents with high heats of absorption. This effect was also confirmed in the results with a generic solvent, reported in the next section. Solvents with high heats of absorption take advantage of the temperature swing. The relative vapor pressure of CO2 and water changes with temperature. This change is greater with solvents with high heats of absorption as shown in
The reboiler duty required for stripping can be approximated as the sum of three terms: the heat required to desorb the CO2, that required to generate the water vapor at the top of the column, and the sensible heat requirement, as shown by the following equation:
Generic solvent modeling. The following three-parameter expression for the vapor-liquid equilibrium was used to model generic solvents:
The constant b was set to 24.76, while the constant a was varied. The value of the constant a, used in the above equation for the generic solvents is shown in
The results show that at 160 kPa, the optimum generic solvent is one with a heat of absorption of 126 kJ/gmol CO2 which is greater than 7 m MEA (80-100 kJ/gmol CO2). At 30 kPa, the optimum generic solvent is one with a heat of absorption 80 kJ/gmol CO2 (about that of 7 m MEA). For solvents with ΔHabs<60 kJ/gmol CO2, stripping at 30 kPa is more attractive than stripping at 160 kPa.
Predicted Performance of Alternative Configurations.
The characteristics of the matrix (265/160 kPa) and 160 kPa strippers for MEA are shown in
The characteristics of the vacuum, and the vacuum internal exchange strippers are in
Effect of Heat of Absorption. From
Effect of Capacity and Mass-Transfer Rates. The capacity of a solvent is defined as the amount of CO2 a solvent can absorb over a given range of loading or partial pressure. This reflects the vapor-liquid equilibrium characteristics of a solvent. A high-capacity solvent can absorb or desorb more CO2 than one with a low-capacity. Some solvents can also achieve richer loading because they have greater rates of CO2 absorption.
The two MEA solvents also have similar heats of absorption. MEA/PZ represented by 11.4 m MEA has a higher capacity and rich loading than 7 m MEA. MEA/PZ offers 13% energy savings over 7 m MEA, with the matrix stripper operated with a 160 kPa reboiler temperature.
On the other hand, KS-1 is never quite as good as MDEA/PZ, even though it has a slightly greater heat of absorption and a slightly greater capacity. The primary difference in this case may be the rich PCO2. Because it is assumed to be a slower reacting hindered amine, KS-1 was assigned a value of 5 kPa compared to 7.5 kPa for MDEA/PZ.
Insight into Stripper Operation. McCabe-Thiele plots provide insight into stripping phenomena.
The McCabe-Thiele plot for 7 m MEA with the matrix (265/160 kPa) configuration is shown in
Concentrated, aqueous solutions of PZ have shown promise for improved solvent performance in absorption/stripping systems for CO2 capture. For 8 m PZ, a CO2 loading of approximately 0.25 mol CO2/mol alkalinity is required to maintain a liquid solution without precipitation at room temperature (20° C.). Additionally, the solubility of PZ at 20° C. is approximately 14 wt % PZ, or 1.9 m PZ. The volatility of 8 m PZ systems was found to be between 10.2 and 18.7 ppm PZ at 40° C., which is comparable to 7 m MEA solutions.
Oxidative degradation of concentrated PZ has been shown to be four times slower than 7 m MEA in the presence of the combination of Fe2+, Cr3+, and Ni2+ and Fe2+ and V4+. In the presence of copper-based corrosion inhibitors, oxidative degradation is an issue but can be drastically reduced with the use of Inhibitor “A”. Concentrated PZ is resistant to thermal degradation up to 150° C. but does degrade at 175° C., losing 32% of the PZ over 2 weeks. The resistance of PZ to thermal degradation allows for the possibility of higher pressure strippers to improve energy performance.
Kinetic measurements have shown that the rate of CO2 absorption into 8 m PZ is more than twice that of 7 m MEA at 40° C. and nearly double at 60° C. The working capacity of an 8 m PZ solution is 0.73 mol CO2/(kg PZ+H2O), nearly double that of 7 m MEA. Initial modeling of the stripper section indicate that the equivalent work required for stripping of an 8 m PZ solution will be approximately 5-10% lower than that of 7 m MEA.
The rapid rate of CO2 absorption, low degradation rate, and low predicted equivalent work indicate that 8 m PZ solutions are an attractive option for CO2 capture in absorption/stripping systems.
Effect on Power Plant Output and Process Improvement. The addition of an absorption/stripping system to a power plant will reduce the plant efficiency by reducing the net power produced from the plant, since steam is withdrawn from the plant to drive the reboiler and electrical power is used to operate compressors, blowers etc. Based on previous process analysis and economic studies, the net power output of a 500 MW power plant is about 150 kJ/gmol CO2 with 90% CO2 removal. Different separation techniques are compared by separation and compression work in
The total equivalent work for reversible isothermal separation to 100 kPa and 313 K, and subsequent compression to 10 MPa, is 18.1 kJ/gmol CO2. This is the theoretical minimum work for separation and compression to 10 MPa. This constitutes about 12% of the power plant output. If reversible, isothermal separation to 100 kPA and 313 K is combined with a real, five-stage, intercooled compressor with 75% adiabatic efficiency, the total equivalent work is 24.1 kJ/gmol CO2 (16% of the power plant output). If an isothermal separation, such as a perfect membrane is used to produce CO2 at the partial pressure in the flue gas at 313 K, and its pressure is increased by a real five-stage, intercooled compressor with 75% adiabatic efficiency, the total equivalent work is 28.4 kJ/gmol CO2.
The best solvent and process configuration is the matrix (295/160 kPa) with MDEA/PZ. This consumes 26.2 kJ/gmol CO2 (18% of the net output from a 500 MW power plant with 90% CO2 capture). This best case offers 22% energy savings over the current industrial baseline (7 m MEA, ΔT=10 K, 160 kPa), and 15% savings over the improved baseline (7 m MEA, ΔT=5 K, 160 kPa). This best case requires 2.1 kJ/gmol CO2 more work than the theoretical minimum with real compressors. Because this analysis is attempting to account for the use of heat and work, it is sensitive to the efficiency selected (75%) for the conversion of steam heat to work.
Multistage Flash Stripper Configuration.
The performance of the multistage flash was estimated for 8 m piperazine. The temperature and CO2 loading of each stage was specified. The basis for the calculation was 0.5 moles/s piperazine (1 equivalent/s piperazine) in the liquid solvent being regenerated.
At each stage the equilibrium CO2 partial pressure (P*CO2, Pa) was calculated as a function of temperature (T,K) and loading (α, moles CO2/mole alkalinity) from an empirical correlation of the available data:
Table 5 gives the estimated total work required to regenerate 8 m PZ at a rich loading of 0.39 moles/mole alkalinity to a lean loading of 0.30. The calculation assumes that lean solution from the last stage of the multistage flash is exchanged to give a rich solution with a feed T that is 5° C. less than the hot lean solution. The vapor produced from each stage is cooled to 40° C. to condense water, then compressed to 150 bar in a multistage compressor intercooled to 40° C.
The least total work, 30.1 kJ.mole CO2, is achieved by a three stage flash with each stage heated to 423 K by steam condensing at 433 K. The pressure of the three stages is respectively, 15, 10, and 7 atm. The CO2 loading at each stage is respectively, 0.36, 0.33, and 0.30 moles/mole alkalinity. With a two stage flash at 423 K, the work increases to 31.0 kJ.mole CO2.
With a simpler three stage flash that heats only at the first stage, the stage T is respectively, 423, 416, and 406. The latter stages operate at lower pressure to provide vapor without additional heating. The stage P is respectively, 15, 8, and 4 atm.
If the three stage flash is operated at 373 K, typical of conventional simple stripper designs, the total work is substantially greater, 41.2 kJ/mole. Furthermore the compressor must be sized to deal with a larger volume of gas at the respective stage pressure, 1.7, 1.3, 1.1 atm, resulting in greater capital cost as well.
Therefore, the present invention is well adapted to attain the ends and advantages mentioned as well as those that are inherent therein. While numerous changes may be made by those skilled in the art, such changes are encompassed within the spirit of this invention as illustrated, in part, by the appended claims.